Process for the catalytic oxidation of hydrocarbons

ABSTRACT

A process for the production of an olefin comprising partially combusting in a reaction zone a mixture of a hydrocabon and an oxygen-containing gas in the presence of a catalyst which is capable of supporting combustion beyond the fuel rich limit of flammability to produce the olefin. Prior the partial combustion, the mixture of the hydrocabon and the oxygen-containg gas is passed through a heat excahanger, and the heat exchanger provides a pressure drop.

The present invention relates, in general, to the catalytic oxidation ofhydrocarbons and, in particular, to the catalytic oxidativedehydrogenation of hydrocarbons to produce olefins.

BACKGROUND OF THE INVENTION

Processes for the catalytic oxidation of hydrocarbons are well known,for example, the oxidation of methane to produce syngas; the oxidationof ethylene to produce ethylene oxide; the oxidation of ethylene andacetic acid to produce vinyl acetate; the production of maleic anhydridefrom the oxidation of butene, butane or benzene; the production ofphthalic anhydride by the oxidation of naphthalene or o-xylene; theammoxidation of propane to acrylonitrile.

The catalytic oxidative dehydrogenation of hydrocarbons is a knownprocess for the production of olefins. An example of such a process isdescribed in EP-A-0 332 289. In this process, a hydrocarbon and anoxygen-containing gas are contacted with a catalyst, which is capable ofsupporting combustion beyond the fuel rich limit of flammability. Thehydrocarbon is partially combusted, and the heat produced is used todrive the dehydrogenation of the hydrocarbon feed into olefins.Optionally, a hydrogen co-feed is also burned, and the heat produced bythis combustion reaction is used to drive the dehydrogenation of thehydrocarbon.

Generally, in a catalytic oxidative dehydrogenation process, thereactants (hydrocarbon and an oxygen-containing gas) are passed over thecatalyst directly to produce olefin product. Typically, the hydrocarbonis a saturated hydrocarbon such as a C₂-C₁₀ saturated hydrocarbon, suchas ethane or a mixture of saturated hydrocarbons such as a mixture ofC₂-C₁₀ hydrocarbons, naphtha or gas oil. The hydrocarbon may be gaseousor liquid at ambient temperature and pressure but is typically gaseous.

It is desirable in processes for the catalytic oxidation of hydrocarbonsto have effective mixing of the reactants and/or uniform velocityprofile of the reactant mixture prior to contact with the catalyst.However, although the reactants (hydrocarbon and oxygen-containing gas)may be pre-mixed prior to being introduced into the reactor, thevelocity profile of the reactant mixture is often still nonuniform. Sucha non-uniform velocity profile can lead to an unstable reaction. Forexample, high velocity gas mixtures over only part of the catalyst canlead to a decrease in selectivity. Where the oxidation reaction is onewhich is carried out close to its flammable limit, such as in thecatalytic oxidative dehydrogenation of hydrocarbons, low velocity gasmixtures over only part of the catalyst can result in flash-backs.

Moreover, the heat generated by the oxidation reaction is generally notdistributed evenly, reducing the efficiency of the oxidation process andresulting in potential loss of product.

SUMMARY OF THE INVENTION

We have now found that the above disadvantages may be alleviated bypassing the reactants (hydrocarbon and an oxygen-containing gas) througha heat exchanger prior to contacting the reactants with the catalyst.

Accordingly, the present invention provides a process for the oxidationof a hydrocarbon said process comprising:

partially oxidising in a reaction zone, a mixture comprising ahydrocarbon and an oxygen-containing gas in the presence of a catalystwhich is capable of supporting oxidation of the hydrocarbon, whereinprior to said partial oxidation, said mixture comprising the hydrocarbonand the oxygen-containing gas is passed through a heat exchanger.

The heat exchanger employed in the present process provides a pressuredrop, which acts to even the velocity profile of the reactant mixture.This ensures that the reactants subsequently flow uniformly over thecatalyst, improving the efficiency of the oxidation reaction.

Suitable heat exchangers include, but are not limited to, shell and tubeor plate fin heat exchangers. The velocity profile of the reactantmixture will generally be dependent on the distance between adjacent gasoutlets of the heat exchanger. Thus, an improved reactant mixturevelocity profile may be achieved by minimising the distance betweenadjacent gas outlets. Suitably, the adjacent gas outlets may be severalmillimeters apart, for example in the range 1 to 10 mm, preferably, 1 to5 mm, especially 2 to 5 mm.

In a preferred embodiment of the present invention the heat exchanger isa compact heat exchanger. Compact heat exchangers generally comprise aplurality of channels, which ensure that fluid (gas or liquid) enteringthe channels develops fully into a substantially uniform velocityprofile within a relatively short distance of exiting the channeloutlets. Preferably, the fluid develops fully into a substantiallyuniform velocity profile within a distance of less than 100 mm fromexiting the channel outlets. More preferably, a substantially uniformvelocity profile may be achieved within a distance of less than 60 mm,even more preferably, less than 40 mm, and most preferably, less than 25mm and especially less than 15 mm from exiting the channel outlets.

The compact heat exchanger may be made by diffusion bonding a pluralityof metal plates together to form a stack. Grooves may be etched orotherwise formed into the surface of each plate, such that the resultingstructure defines a plurality of channels. Preferably, the channels ofthe compact heat exchanger are parallel to one another. The channels maybe of any suitable shape such as linear or z-shaped. The cross-sectionof each channel may be any suitable shape. Preferably, each channel isof the same shape. Preferably, each channel is semi-circular. Where thechannels are semi-circular, the radius of each channel may measure 0.1to 2.5 mm, preferably 0.25 to 4.5 mm, and most preferably, 0.5 to 1 mm.

The compact heat exchanger may be any suitable shape in cross section,for example, circular, rectangular or square. Preferably, the compactheat exchanger is square in cross section. Where a compact heatexchanger of square cross section is employed, each side of the squaremay be in the range 100 to 3000 mm, preferably, in the range 400 and2000 mm, more preferably, in the range 1200 and 1800 mm. Similarly,where a circular compact heat exchanger is employed, the diameter of theexchanger may be in the range 100 to 3000 mm, preferably, in the range400 and 2000 mm, more preferably, in the range 1200 and 1800 mm.

The reactant mixture may be preheated and, preferably, is preheated,prior to entry into the reaction zone. Generally, the reactant mixtureis preheated to temperatures below the autoignition temperature of thereactant mixture. Typically, in the catalytic oxidative dehydrogenationof hydrocarbons, preheat temperatures of up to 200° C. but below theautoignition temperature of the reactants are employed. Preferably, theheat exchanger preheats the reactant mixture to a temperature of up to30° C. below the autoignition temperature of the mixture.

Advantageously, the use of the heat exchanger may allow the reactantmixture to be heated to high-preheat temperatures such as temperaturesat or above the autoignition temperature of the reactant mixture. Thebenefits of using preheat temperatures above the autoignitiontemperature of the reactant mixture will vary depending on theparticular catalytic oxidation process employed. Where the process isthe catalytic oxidative dehydrogenation of hydrocarbon to produceolefin, the use of high pre-heat temperatures is beneficial in that lessoxygen reactant is required which leads to economic savings. Inaddition, in the catalytic oxidative dehydrogenation of hydrocarbon toproduce olefin, the use of high preheat temperatures can result inimproved selectivity to olefin product. It should be understood that theautoignition temperature of a reactant mixture is dependent on pressureas well as the feed composition: it is not an absolute value. Typically,in the catalytic oxidative dehydrogenation of hydrocarbon to produceolefin, where the hydrocarbon feed is ethane at a pressure of 2atmospheres, a preheat temperature of up to 450° C. may be used.

It should be understood that the residence time of the reactant mixturein the heat exchanger will be dependent upon a number of factors such asthe length of the heat exchanger tubes and the velocity of the reactantmixture through the heat exchanger. However, where the reactant mixtureis heated to above its autoignition temperature premature reaction ofthe reactants may occur. To avoid such premature reaction the heatexchanger should preferably provide a residence time of the reactantmixture which is less than the ignition delay time for the temperatureof the reactant mixture. Typically, the residence time may be in therange 10 to 1000 milliseconds, preferably 10 to 100 milliseconds.

In use, the heat exchanger employed in the present invention may beheated by any suitable means. Suitably, the heat exchanger may beindirectly heated by the heat generated by the partial combustion of thereactants. It may also be possible to supplement the heat generated bythe partial oxidation reaction by heating the heat exchanger by externalmeans. For example, the whole or part of the heat exchanger may beheated electrically. Alternatively or additionally, a hot fluid, such assteam, may be passed through the heat exchanger. In certaincircumstances, for example to prevent flashback into the heat exchanger,it may be desirable to cool the whole or part of the heat exchanger.This may be achieved by passing a cooling fluid such as cooling waterthrough the exchanger.

Although the hydrocarbon and oxygen-containing gas may be pre-mixedprior to being introduced into a reaction zone, the extent of mixing ofthe gaseous reactants is often still uneven. Such uneven mixing of thereactants can lead to an unstable reaction.

Advantageously, improved mixing and/or improved velocity profile of thereactants may be achieved by passing the reactant mixture through abaffle zone prior to entry into the heat exchanger.

Thus, in another aspect of the invention, the hydrocarbon andoxygen-containing gas are mixed in at least one baffle zone, prior tobeing passed through the heat exchanger. The baffle zone is defined by ahousing, which contains at least one baffle plate. By baffle plate ismeant any device which serves to correct the flow of reactants therebyimproving mixing and/or velocity profile of the reactants.

The housing may be any suitable shape in cross-section, for example,square, rectangular or circular. A housing having a square cross sectionis preferred. Most preferably, the housing defines a cubic baffle zone.To minimise the volume of the reactant mixture in the housing, it ispreferable to minimise the volume of the baffle zone housing.

The housing may be of any suitable material such as metal.

Suitably, the baffle plate may be solid (non-perforated). Preferably,the baffle plate is of circular cross section. The baffle plate may beof any suitable material such as metal.

Preferably, the baffle plate is disposed substantially at a right angleto the direction of flow of the reactant mixture entering the bafflezone. Such orientation of the plate promotes mixing of the reactants. Ithas been found that where the angle of the plate deviates slightly fromthe perpendicular, for example, by a few degrees, the extent of mixingis inferior to that obtained when the plate is perpendicular to the flowof reactants.

Preferably, the baffle plate is located substantially midway along thelength of the housing such that the distance from the outlet of themixing device to the plate is approximately equal to the distance fromthe plate to the inlets of the heat exchanger channels.

Suitably, where the baffle plate is circular and the housing is ofsquare cross section, the ratio of the length of the housing to thediameter of the baffle plate is in the range 4:1 to 1:1, such as 2:1.

The hydrocarbon and oxygen-containing gas reactants are pre-mixed priorto introduction into the baffle zone. Pre-mixing of the reactants may becarried out by any suitable means. Suitable mixing means include deviceswhich are capable of causing substantial mixing of the reactants priorto entry into the baffle zone such as one or more injectors, forexample, gas injectors. Preferred mixing devices are those which providefor mixing of at least 80% of the reactants prior to entry into thebaffle zone.

Suitably, the hydrocarbon is entrained in the oxygen-containing gasprior to entry into the baffle zone. Preferably, the oxygen-containinggas is injected into the baffle zone by one or more gas injectors.Suitably, the gas injector may be located within a tube or other housingsuch that the oxygen-containing gas stream can be entrained with ahydrocarbon stream entering the tube or housing. The injector may have aflared section around the injector nozzle so as to provide aconstriction in the tube. The hydrocarbon is fed over the flared sectionat lower velocity than the velocity of the oxygen-containing gas. Thedifference in velocities, and the shear between the oxygen containinggas and hydrocarbon streams causes the hydrocarbon to become entrainedin the oxygen-containing gas. Preferably, the velocity ratio of theoxygen-containing gas to hydrocarbon is 20:1 to 2:1, more preferably,3:1 to 8:1. The relative velocities of the reactants and the distancethey have to travel before reaching the baffle zone should be carefullycalculated and monitored, to ensure that the reactants are sufficientlydispersed as they leave the exit port of the injector and enter thebaffle zone.

The reactant mixture entering the baffle zone impinges the baffle plateand is deflected by the baffle plate thereby reducing the momentum ofthe mixture. This reduction in momentum provides a more uniform velocityand pressure profile of the mixture prior to entry into the heatexchanger. The reduction in momentum is enhanced when the baffle plateis positioned in-line with the outlet of the mixing device (such as theexit port of an injector nozzle), that is, the outlet of the mixingdevice is disposed substantially perpendicular to the baffle plate.

Preferably, the dimensions of the baffle plate are such thatsubstantially all of the reactants impinge on the baffle plate. Theratio of the diameter of a circular baffle plate to the diameter of theoutlet of the mixing device is preferably, 1-5:1, more preferably,1-3:1, and most preferably, 2:1.

A plurality of baffle zones may be employed in the process of thepresent invention. Each baffle zone housing may be fed with a singlemixing device or a plurality of mixing devices, preferably a pluralityof mixing devices. Preferably, where a plurality of mixing devices feedsa single baffle zone housing the ratio of baffle plates to mixingdevices is 1:1. However, a baffle zone housing fed by a plurality ofmixing devices may house a single baffle plate.

Preferably, to ensure that the flow rates of reactants through themixing devices are uniform, the outlet of each mixing device hassubstantially the same diameter.

Alternatively, the baffle zone may comprise a housing without a baffleplate. In this instance, the hydrocarbon/oxygen-containing gas mixturemay be fed substantially tangentially into the housing using one or moremixing devices.

Optionally, the hydrocarbon and/or oxygen-containing gas may bepreheated before being introduced into the baffle zone. The temperatureto which the reactants may be preheated, however, is limited by theautoignition temperature of the feed.

Once the reactant mixture has passed through the baffle zone and theheat exchanger, it is contacted with a catalyst which is capable ofsupporting oxidation of the hydrocarbon.

The catalyst which is used in the oxidation reaction of the presentinvention will depend on the specific oxidation process to be employed.For example, where the oxidation process is the oxidation of methane toproduce syngas suitable catalysts include platinum/rhodium or nickelbased catalysts. Suitable catalysts for the oxidation of ethylene toproduce ethylene oxide include silver based catalysts. In the oxidationof ethylene and acetic acid to produce vinyl acetate suitable catalystsinclude palladium based catalysts such as palladium/gold catalysts.Suitable catalysts for the production of maleic anhydride from theoxidation of butene, butane or benzene include vanadium and/ormolybdenum based catalysts. Typically, in the production of phthalicanhydride by the oxidation of naphthalene or o-xylene vanadium basedcatalysts are employed. Suitable catalysts in the ammoxidation ofpropane to acrylonitrile include bismuth based catalysts. These andother suitable catalysts for the afore-mentioned and other hydrocarbonoxidation reactions are known in the art.

Preferably, the oxidation reaction is carried out in a fixed bedreactor.

Any suitable hydrocarbon may be employed, for example, C₁ to C₆hydrocarbons, such as, C₁ to C₆ alkanes, or C₂ to C₆ olefins. The C₁ toC₆ hydrocarbon may be linear, branched or cyclic. Aromatic hydrocarbonssuch as benzene and naphthalene may also be employed.

Any suitable oxygen-containing gas may be employed, for examplemolecular oxygen or air.

Suitably, the oxidation reaction of the present invention is anyhydrocarbon oxidation reaction which may be carried out in a fixed bedreactor. Suitably, the oxidation reaction reaction may be the oxidationof ethylene to ethylene oxide, the oxidation of ethylene and acetic acidto vinyl acetate, the oxidation of napthalene to phthalic anhydride, theoxidation of ortho-xylene to phthalic anhydride, the ammoxidation ofpropane to acrylonitrile, the oxidation of gaseous paraffinichydrocarbons, such as methane, to syngas, the oxidation of C₄hydrocarbon, such as butane and/or butene to maleic anhydride and theoxidation of benzene to maleic anhydride.

The oxidation reaction conditions such as temperature and pressure willdepend on the specific oxidation reaction employed. Suitable reactionconditions are known in the art.

In a preferred embodiment of the present invention there is provided aprocess for the catalytic oxidative dehydrogenation of hydrocarbons toproduce olefins such as ethylene.

Accordingly, the present invention provides a process for the productionof an olefin said process comprising:

partially combusting in a reaction zone, a mixture of a hydrocarbon andan oxygen-containing gas in the presence of a catalyst which is capableof supporting combustion beyond the fuel rich limit of flammability toproduce the olefin, wherein prior to said partial combustion, saidmixture of the hydrocarbon and the oxygen-containing gas is passedthrough a heat exchanger.

The process comprises contacting a hydrocarbon or a mixture ofhydrocarbons and an oxygen-containing gas with a catalyst undercatalytic oxidative dehydrogenation reaction conditions to produce theolefin.

The hydrocarbon may be any hydrocarbon which can be converted to anolefin, preferably a mono-olefin, under the catalytic oxidativedehydrogenation reaction conditions employed.

The process of the present invention may be used to convert both liquidand gaseous hydrocarbons into olefins. Suitable liquid hydrocarbonsinclude naphtha and gas oils. Preferably, however, gaseous hydrocarbonssuch as ethane, propane, butane and mixtures thereof are employed.

Any suitable oxygen-containing gas may be employed, for examplemolecular oxygen, air or molecular oxygen diluted with an unreactive gassuch as nitrogen, argon, carbon dioxide or helium.

Any molar ratio of hydrocarbon to oxygen is suitable provided thedesired olefin is produced in the process of the present invention. Thepreferred stoichiometric ratio of hydrocarbon to oxygen is 5 to 16,preferably, 5 to 13.5 times, preferably, 6 to 10 times thestoichiometric ratio of hydrocarbon to oxygen required for completecombustion of the hydrocarbon to carbon dioxide and water.

The hydrocarbon is passed over the catalyst at a gas hourly spacevelocity of greater than 10,000 h⁻¹, preferably above 20,000 h⁻¹ andmost preferably, greater than 100,000 h⁻¹. It will be understood,however, that the optimum gas hourly space velocity will depend upon thepressure and nature of the feed composition.

Advantageously, the hydrocarbon may be pre-heated. The temperature towhich the hydrocarbon, oxygen-containing gas and (optionally) hydrogenmixture may be preheated, however, is limited by the autoignitiontemperature of the feed.

Preferably, hydrogen is co-fed with the hydrocarbon and molecularoxygen-containing gas into the reaction zone. The molar ratio ofhydrogen to oxygen can vary over any operable range provided that thedesired olefin product is produced. Suitably, the molar ratio ofhydrogen to oxygen is in the range 0.2 to 4, preferably, in the range0.5 to 3.

In the presence of the catalyst, hydrogen combusts preferentiallyrelative to the hydrocarbon, thereby increasing the olefin selectivityof the overall process.

In addition, the feed may comprise a diluent such as nitrogen, carbonmonoxide and steam.

The partial combustion reaction may be suitably carried out at acatalyst exit temperature of between 600° C. and 1200° C., preferablybetween 850° C. and 1050° C. and most preferably, between 900° C. and1000° C.

The catalytic oxidative dehydrogenation process may be carried out atatmospheric or superatmospheric pressure. Suitably, the pressure may bewithin the range 0 to 2 bara, preferably, 1.5 to 2 bara, for example,1.6 to 1.8 bara, especially 1.65 bara. Pressures of, for example, 2 to50 bara, may also be suitable.

The catalyst employed in the catalytic oxidative dehydrogenation processis one which is capable of supporting combustion beyond the fuel richlimit of flammability. Suitable catalysts may comprise one or more GroupVIII transition metals. The Group VIII transition metals include iron,cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium andplatinum. Preferably, the Group VIII transition metal is rhodium andmore particularly, platinum and palladium.

Preferably, the Group VIII transition metal is employed in combinationwith a catalyst promoter. The promoter may be a Group IIIA (aluminium,gallium, indium and thallium), IVA (for example, germanium, tin andlead) and/or VA (for example, arsenic and antimony) metal.Alternatively, the promoter may be a transition metal which is differentto the Group VIII transition metal.

Preferred Group IIIA metals include gallium and indium. Preferred GroupIVA metals include germanium, tin and lead. Preferred Group VA metalsinclude antimony. Preferred transition metals include those from GroupsIB, IIB, VIB, VIIB and VIIIB of the Periodic Table. Examples of suitabletransition metals include chromium, molybdenum, tungsten, iron,ruthenium, osmium, cobalt, rhodium, iridium, nickel, platinum, copper,silver, gold, zinc, cadmium and mercury. Preferred transition metalpromoters are molybdenum, rhodium, ruthenium, iridium, platinum, copperand zinc.

Suitably, the catalyst may be Pt/Ga, Pt/In, Pt/Sn, Pt/Ge, Pt/Cu, Pd/Sn,Pd/Ge, Pd/Cu, Pd/Sn, Pd/Ge and Rh/Sn. Of these Pt/Cu and Pt/Sn arepreferred.

In another embodiment of the present invention, the catalyst comprisesplatinum and palladium, and a further metal selected from Group IIIA,Group IVA or the transition metal series of the Periodic Table.Suitably, the catalyst may be Pt/Pd/Cu.

For the avoidance of doubt, the Group VIII transition metal and promotermay be present in any form, for example, as the metal or in the form ofa metal compound, such as a metal oxide.

Suitable catalytic oxidative dehydrogenation catalysts are described inmore detail in EP-A-0 332 289, WO 97/26987, GB 9930598.9 and GB9930597.1, the contents of which are herein disclosed by reference.

The catalyst employed in the present invention may be unsupported, forexample, the catalyst may be in the form of a metal gauze.

Preferably, the catalyst is supported on any suitable support. Suitably,the support is a metal or ceramic support, preferably a ceramic support.Suitable ceramic supports include lithium aluminium silicate (LAS),alumina (α-Al₂O₃), yttria stabilised zirconia, alumina titanate,niascon, and calcium zirconyl phosphate. The support may be wash-coated,for example, with γ-Al₂O₃.

The catalyst may be employed in the form of tiles, preferably, withchamfered edges. Such tiles may be placed adjacent to one another toform a catalyst bed of a desired size. It may be desirable to seal thearea between the bed and the reactor walls to avoid unreacted reactantsfrom flowing through the gap between the bed and the reactor.Preferably, an intumescent sealant such as vermiculite base is employed.Alternatively, the sealant may comprise compressed silica-fibre mats.

Where possible, the heat produced by the partial combustion of thehydrocarbon and oxygen-containing gas hydrogen is recycled, for example,to heat the reactants in the heat exchanger.

Where the cracking reaction is carried out at superatmospheric pressure,the reaction products may be quenched, for example, with water, as theyemerge from the reaction chamber to avoid further reactions takingplace. Quenching may not be necessary for reactions carried out atrelatively low pressure, for example, pressures less than 5 bara.

Any coke produced in the process of the present invention may be removedby mechanical means, or by decoking. Suitable decoking methods aredescribed in EP 0-A-709 446.

BRIEF DESCRIPTION OF THE DRAWING

The invention will now be illustrated by way of example only and withreference to FIG. 1 and to the following examples.

DETAILED DESCRIPTION OF PREFERRED EMBODIMENTS

FIG. 1 represents in schematic form, apparatus suitable for use in theprocess of the present invention.

The apparatus (1) is provided with means for supplying anoxygen-containing gas (2) and a means for supplying a hydrocarbon andoptionally hydrogen (3). Depending on the scale of the process there maybe a single means or several means for the supply of reactants. Themeans for supplying an oxygen-containing gas (2) and a means forsupplying a hydrocarbon and optionally hydrogen (3) are welded to thebaffle zone housing (4).

The apparatus (1) also comprises a baffle zone housing (4) which isprovided with a baffle plate (5). The baffle zone housing (4) is cubicand made of stainless steel. The baffle plate (5) is solid(non-perforated) and made of monel. The ratio of the length of thebaffle zone housing to the diameter of the baffle plate is 2:1. Thebaffle plate (5) is disposed substantially perpendicular to theoutlet(s) of the reactant supply means and approximately midway alongthe length of the baffle zone housing (4). The ratio of the diameter ofthe baffle plate to the diameter of the exit port of the nozzle of theinjector (8) is 2:1. Depending on the scale of the process, the bafflezone may comprise either a single baffle zone or several baffle zones.Each baffle zone may comprise a single or several baffle plates.

The baffle zone housing (4) is bolted to a heat exchanger (6).

The apparatus (1) further comprises a compact heat exchanger (6) ofsquare cross-section and having substantially z-shaped channels ofsemi-circular cross-section and 1 mm in radius. The heat exchanger (6)is bolted to the reactor (7). Suitable compact heat exchangers aremanufactured by Heatric Limited.

The apparatus (1) also comprises a reactor (7). The cross-sectionalshape of the reactor is square, so as to substantially match thecross-sectional shape of the heat exchanger (6). To minimise heatlosses, the reactor is insulated both internally and externally.

In use, the reactor (7) is provided with at least one catalyst which iscapable of supporting combustion beyond the fuel rich limit offlammability. Suitably, the catalyst is a supported catalyst. Thecatalyst is placed below an alumina heat shield (not shown). Anoxygen-containing gas is fed to the apparatus (1) from a supply (2)suitably via a gas injector (8). A gaseous hydrocarbon feedstocksuitably comprising ethane and optionally hydrogen is fed to theapparatus (1) from supply (3) suitably via a gas injector (8). The ratioof the velocity of the oxygen containing gas to the velocity of thehydrocarbon feed is 5:1. The oxygen-containing gas, hydrocarbon andoptionally hydrogen mix in the injector (8) to form a reactant mixtureprior to exiting the outlet of the injector (8) into baffle zone (4). Onentering baffle zone (4) the reactant mixture impinges the baffle plate(5) and is deflected into the channels of the compact heat exchanger(6). The compact heat exchanger (6) is heated by steam. On passagethrough the channels of the compact heat exchanger (6) the reactantmixture is preheated to the required temperature. On exiting thechannels of the compact heat exchanger (6) the reactant mixture passesinto reactor (7) whereupon it contacts the catalyst. In the reactor (7)at least a portion of the hydrocarbon feed combusts to produce water andcarbon oxides. The hydrogen co-feed, if present, combusts to producewater. Both of these combustion reactions are exothermic and the heatthereby produced is used to drive the dehydrogenation of hydrocarbon toolefin product.

General Ethane Catalytic Oxidative Dehydrogenation Reaction Method

Typically, a platinum-copper catalyst comprising nominal loadings of 3wt % platinum and 1 wt % copper was supported on 99.5% alumina foam (ofporosity 45 pores per inch) (Vesuvius Hi-Tech Ceramics Inc) to provide acatalyst bed sized to match the cross sectional area of heat exchanger,60 mm in depth and a volume of 1008 cm³. The supported catalyst was thenloaded into a fixed bed reactor. An alumina heat shield was positionedabove the catalyst. The catalyst was heated under nitrogen to 200° C. Aflow of nitrogen was maintained immediately below the catalyst to ensurea non-flammable atmosphere until reaction was established on thecatalyst.

Flows of ethane, hydrogen and oxygen were then introduced into anapparatus as described for FIG. 1. The flows of ethane, hydrogen andoxygen were gradually increased over a 1 hour period to provide a gasvelocity at the upstream catalyst face of 4.35 m/s at the feedtemperature with the following feed mass ratios to oxygen:ethane:oxygen1.84 and hydrogen:oxygen 0.12. During this period the catalyst exittemperature rose steadily and then stabilized at 850° C.

After the 1 hour equilibration period the nitrogen flow was reduced to 0kg/h over a 2 hour period. The ethane, hydrogen and oxygen flows wereadjusted to achieve target flows.

The reactant mixture was heated to the required pre-heat temperature bythe compact heat exchanger.

The catalytic oxidative dehydrogenation reaction was carried out at apressure of 1.8 bara.

The product composition was analysed by gas chromatography. The feed andproduct flow rates were determined by coriolis type flow meters.

From analysis of the feed and product flow rates and compositions theethane conversion and selectivity to ethylene was calculated.

The velocity profile and the degree of mixing of the reactants may bedetermined by techniques such as computational fluid dynamic modelling(CFD modelling) and physical flow models. Alternatively and/oradditionally, thermocouples may be located below the catalyst to confirmuniform reaction across the catalyst. Uniform reaction will be obtainedwhen the velocity profile and the extent of mixing of the reactantmixture fed to the catalyst is uniform.

EXAMPLES Example 1

In this Example, ethane, oxygen and hydrogen were utilised as feed inthe general reaction method above. The reactant mixture was pre-heatedto a temperature below the autoignition temperature of the mixture. Theresults are shown in Table 1 below.

Example 2

The data in this Example was obtained by thermochemical analysis of thedata in Example 1. The preheat temperature of the reactant mixture wasset at a temperature above its autoignition temperature. The results areshown in Table 1 below.

TABLE 1 Example 1 Example 2 Preheat Temperature ° C. 195 450Ethane:Oxygen ratio 2.43 3.08 Hydrogen:Oxygen ratio 0.12 0.12 Feedvelocity at catalyst 5.47 7.43 face at preheat temperature m/s Ethaneconversion % 69.14 68.71 Ethylene selectivity 72.29 73.74 (g/100 g C2converted) Carbon monoxide 15.28 12.13 selectivity (g/100 g C2converted) Carbon dioxide selectivity 2.01 1.59 (g/100 g C2 converted)Methane selectivity 8.57 8.74 (g/100 g C2 converted) Acetyleneselectivity 0.59 0.60 (g/100 g C2 converted) C3 selectivity 2.16 2.19(g/100 g C2 converted) C4 selectivity 2.0 2.1 (g/100 g C2 converted) C5+selectivity 0.38 0.38 (g/100 g C2 converted)

It is expected and believed that in the absence of the heat exchangerand/or bafie zone, the conversion of ethane and/or selectivity toethylene would be inferior to the results obtained above.

Example 2 demonstrates that the use of a preheat temperature above theautoignition temperature of the reactant mixture results in improvedselectivity to olefin product and a reduction in oxygen usage comparedto the use of a preheat temperature below the autoignition temperatureof the reactant mixture (Example 1).

Example 3

Using the apparatus as described in FIG. 1, the distribution of oxygenin an oxygen/ethane gas mixture (ratio of oxygen to ethane, 1:2.4) fedinto the baffle zone (4) by the injector (8), passing through the bafflezone (4) and into the compact heat exchanger (6) was analysed by 3D CFDmodelling using Fluent™ software. In this Example the baffle plate (5)was disposed perpendicularly i.e at 90 degrees to the oxygen/ethane gasflow.

The CFD modelling results demonstrated that the oxygen and ethanepartially mix in the injector (8) prior to entry into the baffle zone(4). After deflection by the baffle plate (5) the oxygen and ethanegases appear to be substantially mixed (approximately 99% mixed)

No further mixing of the gases was observed after exit from the bafflezone (4).

Example 4

Example 3 was repeated except that the baffle plate (5) was placed at anangle to the oxygen/hydrocarbon gas flow of 2.6 degrees.

The CFD modelling results show that the mixing of the oxygen and ethanegases after deflection by the baffle plate (5) was not as good(approximately 96% mixed) as that obtained in Example 4 (where baffleplate was perpendicular to the gas flow).

Examples 3 and 4 clearly demonstrate that the mixing of oxygen andethane is not complete at the baffle plate. Thus, in the absence of abaffle zone there would be a range of ethane/oxygen gas mixturescontacting the catalyst resulting in a less efficient auto-thermalcracking process.

The CFD modelling also demonstrated that in both Examples 3 and 4 themomentum of the injected reactant mixture was reduced by the baffleplate (5) resulting, in each case, in an improved, that is a moreuniform velocity profile of the reactant mixture prior to entry into theheat exchanger (6).

1. A process for the production of an olefin, said process composing: partially combusting in a reaction zone a mixture of a hydrocarbon and an oxygen-containing gas in the presence of a catalyst which is capable of supporting combustion beyond the fuel rich limit of flammability to produce the olefin, wherein prior to said partial combustion, said mixture of the hydrocarbon and the oxygen-containing gas is passed through a heat exchanger, and wherein the heat exchanger provides a pressure drop.
 2. A process according to claim 1 wherein the heat exchanger has gas outlets and the distance between adjacent gas outlets is in the range 1 to 10 mm.
 3. A process according to claim 1 wherein the heat exchanger provides a residence time in the range 10 to 1000 milliseconds.
 4. A process according to claim 1 wherein the heat exchanger is a compact heat exchanger.
 5. A process according to claim 4 wherein the fluid exiting the channel outlets of the compact heat exchanger develops into a substantially uniform velocity profile within a distance of less than 100 mm from the channel outlets.
 6. A process according to claim 5 wherein a substantially uniform velocity profile is achieved within a distance of less than 25 mm from the channel outlets.
 7. A process according to claim 4 wherein the compact heat exchanger has a square cross section.
 8. Process according to claim 7 wherein each side of the square is in the range 100 to 3000 mm.
 9. A process according to claim 1 wherein the heat exchanger preheats the mixture of the hydrocarbon and oxygen-containing gas to a temperature below the autoignition temperature of the mixture.
 10. A process according to claim 1 wherein the heat exchanger preheats the mixture of the hydrocarbon and oxygen-containing gas to a temperature at or above the autoignition temperature of the mixture.
 11. A process according to claim 1 wherein the hydrocarbon and oxygen-containing gas are mixed in at least one baffle zone prior to being passed through the heat exchanger.
 12. A process according to claim 11 wherein the baffle zone comprises a housing and at least one baffle plate is contained within the housing.
 13. A process according to claim 12 wherein the baffle plate is of circular cross section.
 14. A process according to claim 12 wherein the baffle plate is disposed substantially perpendicularly to the direction of flow of the hydrocarbon and oxygen-containing gas mixture.
 15. A process according to claim 12 wherein the baffle plate is located substantially midway along the length of the housing.
 16. A process according to claim 12 wherein the housing defines a cubic baffle zone.
 17. A process according to claim 16 wherein the baffle plate is of circular cross section and the ratio of the length of the housing to the diameter of the baffle plate is in the range 4:1 to 1:1.
 18. A process according to claim 11 wherein the hydrocarbon and oxygen-containing gas mixture is pre-mixed in at least one mixing device prior to entry into the baffle zone.
 19. A process according to claim 18 wherein the baffle zone comprises a circular baffle plate and the ratio of the diameter of the baffle plate to the diameter of the outlet of the mixing device is 1-5:1.
 20. A process according to claim 18 wherein a plurality of mixing devices feed a single baffle zone housing which comprises a plurality of baffle plates.
 21. A process according to claim 20 wherein the ratio of baffle plates to mixing devices is 1:1.
 22. A process according to claim 11 wherein the baffle zone comprises a housing into which the hydrocarbon/oxygen-containing gas mixture is fed substantially tangentially using one or more mixing devices.
 23. A process according to claim 1 wherein the catalyst comprises one or more Group VIII transition metals.
 24. A process according to claim 1 wherein the stoichiometric ratio of hydrocarbon to oxygen is 5 to 16 times the stoichiometric ratio of hydrocarbon to oxygen required for complete combustion of the hydrocarbon to carbon dioxide and water.
 25. A process according to claim 1 wherein the partial combustion is carried out in the presence of hydrogen.
 26. A process according to claim 1 wherein the partial combustion is carried out at a catalyst exit temperature in the range 600° C. to 1200° C.
 27. A process according to claim 1 wherein the partial combustion is carried out at a pressure in the range 0 to 2 bara. 